Hydrogenation and dehydrogenation processes and catalysts therefor

ABSTRACT

A process for hydrogenating unsaturations in petrochemical feedstocks, the process comprising contacting the petrochemical feedstock, including at least one component having unsaturations, and hydrogen with a catalyst comprising at least one Group Ia, Ib, IIb, VIb, VIIb or VIII metal on a support of an alkaline earth metal silicate having a surface area of at least 30 m 2 /g at a temperature of from 0 to 250° C. and a pressure of from 3 to 50 barg.

The present invention relates to a process for hydrogenating unsaturatedpetrochemical feedstocks, in particular a process for the selectivehydrogenation of such feedstocks. The present invention also relates toa process for dehydrogenating petrochemical feedstocks. The presentinvention also relates to a catalyst, in particular a catalyst for usein such a process.

There are a number of known processes for the hydrogenation ofunsaturated hydrocarbons. For most applications, the hydrogenation mustbe carried out in a selective manner, i.e. some unsaturated hydrocarbonshave to be hydrogenated while other unsaturated hydrocarbons may not behydrogenated. Among the pure hydrocarbons, three kinds of unsaturatedhydrocarbons can be considered: (1) multiple unsaturated hydrocarbonsare alkynes with triple bonds, diolefins with two double bonds or evenmultiple olefins with more double bonds; (2) unsaturated hydrocarbonswith only one double bond; and (3) aromatic unsaturated hydrocarbons,having an aromatic nucleus. Selective hydrogenation means that only oneor two of the three unsaturates are reduced. Very important industrialapplications of hydrogenation are: (1) removal of impurities from steamcracker product streams, for example selective hydrogenation of multipleunsaturated hydrocarbons in olefin rich streams, or selectivehydrogenation of multiple unsaturates and unsaturated hydrocarbons fromaromatic rich streams, and (2) hydrogenation of macromolecules, forexample hydrocarbon solvents and base oils, polyalpha-olefins and evenresins, polymers and copolymers.

Heterogeneous hydrogenation catalysts contain an active metal compoundon a carrier.

Among the active metal compounds are Group IIb, VIb, VIIb and VIIIelement compounds. They can be in the metallic state, in an oxidicstate, in a partially reduced oxide state or even in a sulphided orpartially sulphided state. Also metallic Group Ia metals are known to beactive hydrogenation catalysts. The most preferred metals or metalcompounds are those of Pd, Pt, Ni, Rh, Co, Fe, Cu, Ir, Ru, Os, W, Mo andNa or K. All these active hydrogenation catalysts can also exhibitisomerisation activity to some extent. It is known that in particularNa, K, Fe, Pd and Ni metals catalyse double bond migration, while Pt andCu are much less active isomerisation catalysts. Activity andselectivity of selective hydrogenation catalysts can still further beimproved by employing bimetallics or bimetallic compounds. Typicalexamples are CoMo, NiW and NiMo sulphided catalysts used forhydrotreatment. Other examples used for selective hydrogenation are:Cu—Pd, Cu—Ni, Cu—Co, Cu—Pt, Fe—Pd, Co—Pd, Ni—Pd, Pt—Pd, Ag—Pd, Fe—Pt,Ni—Pt, Pt—Sn, Pt—Pb, Pd—Sn, Pd—Pb, Au—Pd and many others.

It is known that activity and selectivity can also be influenced by thecharacteristics of the carrier for the metal compound. The carrier caninfluence the dispersion of the metal or metal compound, the particlesize of the metal of metal compound and the electronic properties of themetal or metal compound.

Known carriers include carbon, alumina, silica, titania, zirconia, saltsof alkaline earth metals and zeolites or molecular sieves. The acid-baseproperties of the carrier can be very important for several reasons. Thecarrier properties can influence the dispersion of the metal or metalcompound, its electronic properties and hence its activity andselectively. Moreover, when the carrier is not completely covered withmetal or metal compound, the remaining acid and/or basic sites mayinfluence the catalytic behaviour of the catalyst. When highlyunsaturated hydrocarbons are to be hydrogenated, they will interactstrongly with acidic carriers whereas they will interact little withbasic carriers. Even when hydrogenation needs high temperature, sitereactions such as acid catalysed isomerisation or cracking can occur. Itis known that addition of basic compounds during hydrogenation or addingbasic metal compounds on the carrier do increase catalytic performance.Known basic carriers are salts of alkaline earth metals. However, theydevelop only very low surface areas.

A number of catalysts for selective hydrogenation of unsaturatedhydrocarbons are available commercially. Such catalysts comprise, forexample palladium on an alumina support, palladium on an activatedcarbon support, nickel tungsten on an alumina support and palladium on abarium sulphate support.

In complete contrast to the hydrogenation process, it is also known todehydrogenate hydrocarbons. For example, it is known in the art thatdehydrogenation catalysts for the dehydrogenation of light paraffinssuch as propane and butane primarily employ supported platinum, nickelor chromium. In such supported platinum catalysts, the platinum ispresent as a metal and is often promoted with tin. Chromium-basedcatalysts contain chromium oxide as the active phase. The nickel-basedcatalysts mainly employ nickel in the form of sulphide which is presenton the support. For these three catalyst types it is known in the artthat the carrier of the active phase has a very important effect oncatalyst performance—the activity, the selectivity and the stability areall influenced by the support. Often, these three kinds of catalysts aresupported on alumina-type carriers which have been modified by theaddition of one or more alkali metal or alkaline earth metal compounds,which tend to moderate the acidity of the alumina in the carrier andhence increase the selectivity and the potential lifetime of thecatalyst. On the other hand, the active metal compound may be supportedon a spinel-like carrier such as MgAl₂O₄ or ZrO₂ which are less acidicthan alumina-type carriers and which exhibit a high thermal stability.The property of high thermal stability is very important, sincedehydrogenation reactions typically require a temperature of from 500 to630° C.

Another catalytic dehydrogenation process is reforming which is a veryimportant refinery application, in which the main goal is todehydrogenate alkyl cycloparaffins into aromatics with co-production ofhydrogen. Conventional reforming catalysts typically comprise platinumsupported on an acidified alumina carrier. The acidic function isrequired when isomerisation and dehydrocyclisation are desired toconvert additional paraffins into isomers and aromatics. There is aninterest in the art to convert only paraffins with at least 6 carbonatoms into aromatic compounds with co-production of hydrogen. It isknown in the art that this may be done over a catalyst comprising abasic zeolite carrier impregnated with platinum. Since there are noacidic sites associated with the basic zeolite carrier, competingreactions such as isomerisation and hydrocracking are suppressed,resulting in a very high selectivity for the production of aromaticcompounds.

However, there is a need in the art for improved dehydrogenationcatalysts, and in particular carriers for such dehydrogenationcatalysts.

A huge variety of naturally occurring and synthetically producedsilicates are known in the art.

U.S. Pat. No. 3,806,585 discloses the production of a hydrous calciumsilicate composed preponderantly of xonotlite in the shape of rodcrystals which is described as having outstanding refractory properties,whereby moulded bodies comprised primarily of xonotlite provide strengthunattained by other inorganic materials. The specification disclosesthat hydrous calcium silicate of the xonotlite type has use inconstruction as a fire proof coating material, as a fire proof moistureretaining material and as a potential filler for plastics and rubberproducts.

U.S. Pat. No. 3,804,652 discloses a method of producing calcium silicateproducts, such as drain pipes and insulating material, to formtobermorite having the empirical formula 5CaO.6SiO₂.5H₂O.

U.S. Pat. No. 3,928,539 discloses a method of producing hydrous calciumsilicates such as xonotlite, tobermorite and the like.

U.S. Pat. No.3,915,725 discloses a process for producing hollowspherical aggregates of xonotlite, which aggregates are employed to formshaped articles.

U.S. Pat. No.4,298,386 discloses the production of globular secondaryparticles of the woolastonite group of calcium silicate crystals,including woolastonite and xonotlite.

U.S. Pat. No.4,689,315 discloses the production of amorphous,approximately spherical silica particles obtained by the acidichydrolysis of an approximately spherical synthetic calcium silicate. Theresultant silica particles, obtained by such acid hydrolysis, aredisclosed as being particularly suitable for use as catalyst support.The starting material may comprise spherical synthetic calcium silicatessuch as xonotlite, tobermorite and/or calcium silicate hydrate, whichare then treated with an aqueous acid having a pH of from 0.6 to 3 toproduce the resultant silica particles for use as a catalyst support.

U.S. Pat. No.4,849,195 discloses synthetic substantially sphericalcrystal aggregates of xonotlite. The aggregates can be mixed with inertparticles, for example to produce thermal insulation products.Alternatively, as for U.S. Pat. No.4,689,315 described above, theaggregates of xonotlite can be used as starting material for acidextraction of calcium atoms in order to obtain silica.

The present invention in one preferred aspect aims to provide animproved method of selectively hydrogenating unsaturated petrochemicalfeedstocks.

Accordingly, the present invention provides a process for hydrogenatingunsaturations in petrochemical feedstocks, the process comprisingcontacting the petrochemical feedstock, including at least one componenthaving unsaturations, and hydrogen with a catalyst comprising at leastone Group Ia, Ib, IIb, VIb, VIIb or VIII metal on a support of analkaline earth metal silicate having a surface area of at least 30 m²/gat a temperature of from 0 to 550° C. and a pressure of from 3 to 150barg.

Preferably, the at least one petrochemical feedstock is passed-over thecatalyst at an LHSV of from 1 to 100 h⁻¹.

The molar ratio of hydrogen to the at least one component havingunsaturations to be selectively hydrogenated may be from 0.7 to 200.

The present invention also provides a process for hydrogenatingunsaturations in petrochemical feedstocks, the process comprisingcontacting the petrochemical feedstock, including at least one componenthaving unsaturations, and hydrogen with a catalyst comprising at leastone Group Ia, Ib, IIb, VIb, VIIb or VIII metal on a support of acrystalline calcium silicate having the chemical compositionCa₆Si₆O₁₇(OH)₂.

The present invention further provides a catalyst comprising at leastone Group Ia, Ib, IIb, VIb, VIIb or VIII metal on a crystalline calciumsilicate support having a surface area of at least 30 m²/g, the supportbeing in the form of substantially spherical particles and pores in theparticles having a diameter of from 100 to 2000 Angstroms.

Preferably, the particles have a mean diameter of from 10 to 200microns.

The catalyst may be used, in accordance with the invention, in a processfor hydrogenating an unsaturated hydrocarbon feedstock or in a processfor dehydrogenating or reforming a hydrocarbon feedstock.

The present invention yet further provides a catalyst comprising a metalon a support comprising a crystalline calcium silicate of molecularformula 6CaO.6SiO₂.H₂O.

Preferably, the metal comprises at least one metal selected from GroupsIa, Ib, IIb, VIb, VIIb and VIII of the periodic table.

The present invention still further provides the use of a crystallinecalcium silicate of molecular formula 6CaO.6SiO₂.H₂O as a catalystsupport.

The present invention is at least partly predicated on the surprisingdiscovery that a basic hydrated crystalline calcium silicate when usedas a catalyst support can yield hydrogenation catalysts for selectivehydrogenation of petrochemical feedstocks having high activity andselectivity. This is all the more surprising since xonotlite-typematerials have been known for a number of years but to the applicant'sknowledge there has been no disclosure or suggestion in the prior art ofusing xonotlite-type materials as catalysts or catalyst carriers.Rather, as disclosed in for example U.S. Pat. No. 4,689,315 as discussedabove, xonotlite has been proposed in the prior art for use as astarting material for the production of silica, where the chemicalcomposition and structure of the xonotlite is destroyed in thepreparation of the silica particles by acid hydrolysis.

The present invention is also at least partly predicated on thesurprising discovery that a basic hydrated crystalline calcium silicatecomprising xonotlite is a is a suitable carrier for dehydrogenation andreforming reactions because at temperatures of up to 650° C., such abasic carrier has high temperature stability, in that the carrierretains its crystallinity and substantially retains its pore volume andsurface area.

Preferred embodiments of the present invention will now be described ingreater detail by way of example only.

The catalyst of the present invention preferably comprises a supportednoble metal catalyst.

The catalyst of the present invention comprises at least one Group Ia,Ib, IIb, VIb, VIIb or VIII metal, such as Pd, Co, Rh, Ru, Ni, Mo, W, Fe,Cu, Na or K or a combination thereof with palladium being particularlypreferred.

The metal or metals may be in the metallic state, in an oxidic state, ina partially reduced oxide state, or in a sulphided or partiallysulphided state. Optionally, bi-metallic metals or bi-metallic compoundsmay be incorporated into the hydrogenation catalyst, such as CoMo, NiW,and NiMo sulphided catalyst for hydro-treatment and, for selectivehydrogenation, Cu—Pd, Cu—Ni, Cu—Co, Cu—Pt, Fe—Pd, Co—Pd, Ni—Pd, Pt—Pd,Ag—Pd, Fe—Pt, Ni—Pt, Pt—Sn, Pt—Pb, Pd—Sn, Pd—Pb and Au—Pd.

The preferred catalyst support is a basic alkaline earth metal silicatewith a very open and accessible pore structure. A most preferredcatalyst support comprises a synthetic crystalline hydrated calciumsilicate having a chemical composition of Ca₆Si₆O₁₇(OH)₂ whichcorresponds to the known mineral xonotlite (having a molecular formula6CaO.6SiO₂.H₂). The catalyst support preferably has a sphericalmorphology with a mean diameter of the spherical particles being from 10to 200 μm. The support has a very open structure comprising an outershell with a very close-crystal structure surrounding an open innerstructure. This may be referred to as an egg shell like structure. Theouter shell is formed of interlocked ribbon-shaped crystals yieldingregular and homogeneous surface properties. The outer shell is providedwith pore openings up to 2000 Angstroms, more preferably from 100 to1000 Angstroms, in diameter, This provides a good pore structure withhigh pore volume.

Preferably, the support has a specific surface area well above 10m^(2/)g, ranging from 30 to 200 m²/g, more preferably from 40 to 90m²/g.

The support material is preferably pH basic. More preferably, thesupport material has a minimum basicity corresponding to a pH of greaterthan 7.5. The pH may be measured when 4 wt % of the support material isimmersed in water.

Generally, a synthetic hydrated calcium silicate is synthesisedhydrothermally under autogeneous pressure. A particularly preferredsynthetic hydrated calcium silicate is available in commerce from thecompany Promat of Ratingen in Germany under the trade name Promaxon D.This material exhibits some basicity due to the presence of calcium, andin a 4% by weight dispersion in water, the pH reaches a value of around10. The specific composition of the preferred synthetic hydrated calciumsilicate is specified in Table 1.

In order to demonstrate the thermal stability of xonotlite, andtherefore the applicability of xonotlite as a carrier fordehydrogenation and reforming reactions, commercial xonotlite Bold underthe trade name Promaxon D was calcined in ambient air at a relativehumidity of about 50% at two different temperatures, namely 650° C. and750° C., each for a period of 24 hours. The initial xonotlite had acrystalline phase Ca₆Si₆O₁₇(OH)₂ with a BET surface area of 51 m²/gramand a pore volume (of less than 100 nanometers) of 0.35 ml/gram. Aftercalcination at 650° C., the carrier retained its crystallinity whichcorresponds to that of xonotlite. Thus after a 24 hour calcination at650° C., the crystalline phase still comprised xonotlite(Ca₆Si₆O₁₇(OH)₂) with a BET surface area of 47.4 m²/gram and a porevolume (less than 100 nanometers) of 0.30 ml/gram. After the calcinationat 750° C., the carrier was transformed into wollastonite (having thecrystalline phase CaSiO₃) by losing one water molecule. This made thecarrier less basic. Furthermore, as a result of calcination at 750° C.the carrier lost much of its pore volume, being reduced to 0.09 ml/gram(for pore sizes of less than 100 nanometers) and the BET surface areawas correspondingly reduced to 38 m²/gram.

These results show that xonotlite has utility as a basic carrier forhigh temperature reactions in the range of from 500 to 650° C., moreparticularly from 500 to 630° C., the typical temperature range fordehydrogenation and reforming reactions. In these temperature ranges thexonotlite retains its basicity, resulting in the carrier being suitablefor incorporation in a catalyst for use in reforming reactions.

The at least one Group Ia, Ib, IIb, VIb, VIIb or VIII metal ispreferably present in an amount of from 0.01 to 10 wt %, more preferablyabout 0.5 wt %, based on the weight of the supported catalyst.

The catalyst is produced by impregnating the at least one Group Ia, Ib,IIb, VIb, VIIb or VIII metal on the alkaline earth metal silicatesupport. Preferably, an incipient wetness impregnation technique isemployed where the pores of the support are filled with a volume ofsolution containing the metal. In this technique, the dried catalyst isimpregnated with a solution of a salt of the at least one Group Ia, Ib,IIb, VIb, VIIb or VIII metal, for example a halide of the metal, inparticular the Group VIII metal chloride. The amount of the metal saltis calculated to provide a desired metal content on the support, forexample a metal content of from 0.01 to 10 wt % based on the weight ofthe supported catalyst, most preferably about 0.5 wt % based on theweight of the supported catalyst. The impregnated solid is dried firstunder vacuum and subsequently at elevated temperature. Finally, theproduct is calcined, for example at a temperature of about 250° C. for aperiod of about 3 hours.

Alternatively an excess of solution is used during the impregnation stepand the solvent is removed by evaporation. Depending on the propertiesof the impregnation solution and the carrier the active metal phase canhave different locations: (1) the metal or metal compound isconcentrated in a thin layer close to the external surface, this may bereferred to as an “egg-shell mode”, (2) the metal or metal compound isconcentrated in a thin layer below the surface, but not penetrating tothe centre, this may be referred to as an “egg-white mode”, (3) themetal or metal compound is concentrated in a small zone near the centreof the particle carrier, this may be referred to as an “egg-yolk mode”,and (4) the metal or metal compound is uniformly distributed throughoutthe particle carrier. The way that the metal precursor will interactwith the carrier depends on the isoelectric point (IEP) which is the pHat which the particle of the carrier in an aqueous solution has no netcharge. At pH's above the IEP, cations will be adsorbed, because thesurface carries a negative charge; below the IEP, only anions will beadsorbed, because the surface carries a positive charge. During thecontact of the impregnating solution and the carrier, ion exchange canalso occur. The impregnating solution may be altered by addingcomplexing agents, which can change the charge of the metal precursor.In another technique, competing ions may be added to improve thespreading of the metal precursor over the carrier.

In alternative embodiments of the catalyst production process, the metalmay be deposited on the support by ion exchange or vapour phasedeposition.

The catalyst of the present invention is a heterogeneous catalyst whichmay be used in a batch wise or continuous process. Preferably, thecatalyst is used in a fixed bed reactor. A most preferred processemploys a continuously operated fixed bed reactor.

In the hydrogenation process, the petrochemical feedstock is contactedbatch-wise or continuously passed over the catalyst at a selectedtemperature and pressure. The temperature is preferably from 0 to 250°C. The total pressure is preferably from 3 to 50 bar. The petrochemicalfeedstock is preferably contacted with the catalyst at a liquid hourlyspace velocity (LHSV) of from 0.1 to 100 h⁻¹, more preferably from 1 to100 h⁻¹.

The hydrogenation conditions vary dependent on the nature of thepetrochemical feedstock and the process of the invention may be employedfor hydrogenating a variety of different unsaturated petrochemicalfeedstocks. Fundamentally, the feedstocks are those to be selectivelyhydrogenated where one of two unsaturates is reduced or one or two ofthree unsaturates are reduced, the unsaturates being selected frommultiple unsaturated hydrocarbons such as alkynes with triple bonds,diolefins with two double bonds or multiple olefins with more doublebonds; unsaturated hydrocarbons with only one double bond; and aromaticunsaturated hydrocarbons having an aromatic nucleus.

In a first preferred aspect, the process is used for selectivehydrogenation of butadiene to butenes in crude C4 streams.

The C4 streams may come from an FCC unit, a visbreaker or a coker, ormay comprise a C4 stream from a steam cracker or a C4 fraction of anethylene plant. The C4 fraction of an ethylene plant contains highconcentrations of butadiene (typically 25 to 75 wt %). It is desirableto hydrogenate such butadiene into butenes for further processing.Moreover, a C4 fraction from which the butadiene has been removed byconversion or extraction may still contain residual butadiene.Typically, in this preferred process the C4 stream containing butadieneis fed over the catalyst together with hydrogen so as to have ahydrogen/butadiene molar ratio of from 1 to 10, under process conditionscomprising an inlet temperature of from 20 to 200° C., a total pressureof from 5 to 50 barg and an LHSV of from 1 to 40 h⁻¹. The reactoreffluent may be recycled in order to control the outlet temperature.Optionally, several reactors in series may be used with intermittentcooling and/or injection of hydrogen for improved control of thehydrogen content in the feedstock.

In a second preferred aspect of the invention the process may beemployed for selective hydrogenation of vinyl-and ethyl acetylenes incrude C4 streams. The C4 streams typically come from steam crackers. TheC4 fraction from an ethylene plant contains, beside the butadienediscussed hereinabove, varying amounts of vinyl acetylene and ethylacetylene. These have to be removed before further processing, such asby extraction or conversion. The feedstock is fed together with hydrogenover the catalyst, there being a hydrogen/butadiene molar ratio of from1 to 10 under process parameters having an inlet temperature of from 0to 100° C., a total pressure of from 3 to 35 barg and an LHSV of from 1to 40 h⁻¹. Again, the reactor effluent may be recycled in order tocontrol the outlet temperature and optionally several reactors in seriesmay be used with intermittent cooling and/or injection of hydrogen forimproved control of the hydrogen content in the feed.

In a third preferred aspect of the invention the process may be employedfor selective hydrogenation of methyl acetylene and propadiene topropylene in C3 streams. The feedstock typically comprises a C3 cut froma steam cracking unit, which most typically is a C3 fraction with highpropylene content which is obtained from an ethylene plant. Thisfraction contains methyl acetylene and propadiene. These compounds haveto be removed for further processing of the propylene. In this aspect ofthe process, the feedstocks is fed together with hydrogen, at ahydrogen/MAPD molar ratio of from 0.7 to 5 (MAPD being the total molarcontent of methyl acetylene and propadiene) under process parameterscomprising an inlet temperature of from 0 to 100° C., a total pressureof from 10 to 50 barg and an LHSV of from 10 to 50 h⁻¹. The reaction maybe carried out in a multi-tubular pseudo-isothermal reactor or in anadiabatic reactor. As for the other preferred aspects, the reactoreffluent may be recycled in order to control the outlet temperature andoptionally several reactors in series may be used with intermittentcooling and/or injection of hydrogen in order to provide better controlof the hydrogen content in the feedstock.

In accordance with a fourth preferred aspect of the invention, theprocess of the invention may be employed for the selective hydrogenationof pyrolysis gasoline, which may also be known in the art as “firststage” hydrogenation of the pyrolysis gasoline. The feedstock comprisespyrolysis gasoline from steam cracking units, coker units orvisbreakers. In accordance with this aspect, diolefins and unsaturatedaromatics are converted into the corresponding olefins and aromatics.The hydrogenated product can be used as a stable gasoline blending feedor can be further hydrotreated for the recovery of aromatics. Thefeedstock is passed over the catalyst together with hydrogen to providea hydrogen/diene molar ratio of from 1 to 10 under the processparameters of an inlet temperature of from 20 to 200° C., a totalpressure of from 5 to 50 barg and an LHSV of 1 to 20 h⁻¹. Again, as forthe other aspects, the reactor effluent may be recycled in order tocontrol the outlet temperature and optionally several reactors in seriesmay be used with intermittent cooling and/or injection of hydrogen inorder to achieve better control of the hydrogen content in thefeedstock.

In a yet fifth preferred aspect of the invention the process of theinvention is employed for selective hydrogenation of gasoline fractions.The feedstock may comprise fractions from pyrolysis gasoline originatingfrom steam cracking units, coker units or visbreakers and light crackednaphthas from FCC units. In this aspect, dienes and acetylenes in thegasoline fractions are selectively removed for the preparation ofethers. The feedstock is passed over the catalyst together with hydrogento provide a hydrogen/diene molar ratio of from 1 to 20, and the processparameters are an inlet temperature of from 20 to 250° C., a totalpressure of from 5 to 50 barg and an LHSV of from 1 to 20 h⁻¹.

In a sixth preferred aspect of the invention, the process of theinvention may be employed for selective hydrogenation of phenylacetylene in crude styrene streams. The feedstock comprises crudestyrene. Crude styrene production by dehydrogenation of ethyl benzene orrecovery from pyrolysis gasoline tends to yield styrene containing smallamounts of phenyl acetylene which has to be removed before furtherprocessing. The styrene is fed together with hydrogen to yield ahydrogen/phenyl acetylene molar ratio of from 1 to 20 over the catalystat an inlet temperature of from 10 to 150° C., a total pressure of from5 to 50 barg and an LHSV of from 10 to 100 h⁻¹.

In a seventh preferred aspect of the invention, the process of theinvention is for selective hydrogenation of olefins in aromatic richfractions. The feedstocks may comprise aromatic rich fractions fromreforming units, from cokers or from steam cracking units. Such aromaticrich fractions need to be treated to extract the aromatics. Before theextraction of the aromatics, the residual bromine index (which reflectsthe olefin content) has to be very low. Any process to reduce the olefincontent by hydrogenation needs to minimize the conversion of thearomatics. Also, a further reduction of the bromine index in almost purearomatic fractions may require a further hydrogenation step which canreplace conventional clay treatment. The feedstock is passed togetherwith hydrogen over the catalyst at a hydrogen/olefins molar ratio offrom 5 to 100 under process parameters comprising an inlet temperatureof from 5 to 250° C., a total pressure of from 5 to 50 barg and an LHSVof from 5 to 50 h⁻¹.

In an eight preferred aspect or the present invention, the process ofthe invention may be employed for selective hydrogenation ofpetrochemical feedstocks in conjunction with a reforming process.

When the crystalline calcium silicate support (such as xonotlite) isused for a dehydrogenation or reforming catalyst, the catalyst, as wellas the support, comprises at least one Group IIb, VIb, VIIb or VIIImetal such as Pd, Co, Rh, Ru, Ni, Mo, W, Fe, Cu or a combinationthereof. The feedstocks for dehydrogenation may typically comprise lightparaffins, such as propane and butane. The feedstocks for reformingreactions may typically comprise normal paraffins and cycloparaffins,such as n-hexane and cyclohexane. The dehydrogenation and reformingreactions may be carried out at a temperature of from 500 to 630° C.

The present invention will now be described with reference to thefollowing non-limiting Example.

EXAMPLE 1 Catalyst Preparation

Extrudates of the hydrated crystalline calcium silicate available incommerce under the Trade name Promaxon D was dried at a temperature of500° C. for a period of 3 hours. The dried support gas then impregnatedwith a solution of palladium chloride (PdCl₂) using a wet impregnationtechnique. In particular, 65.38 g of dried Promaxon D were progressivelycontacted with 38.23 mol of an aqueous palladium chloride solution, theamount of solution being selected so as to correspond to the estimatedabsorption capacity of the dried Promaxon D. The amount of the palladiumsalt was calculated in order to reach a final palladium content in theresultant catalyst of 0.3 wt %. The impregnated solid was dried undervacuum for a period of 36 hours at 25° C. and thereafter dried for aperiod of 16 hours at a temperature of 110° C. Finally, the catalyst wascalcined at a temperature of 400° C. for a period of 3 hours.

Selective Hydrogenation of Pyrolysis Gasoline

An amount of 42.2 g (having a volume of 75 ml) of the activated catalystcomprising 0.3 wt % Pd on the xonotlite carrier was transferred undernitrogen into a laboratory scale continuous trickle bed reactor. Thecatalyst was then reduced under a flowing hydrogen stream at 120° C.Thereafter a pyrolysis gasoline from a steam cracker having thecomposition and properties specified in Table 2, was passed through thereactor at an LHSV of 4.92 h⁻¹ (corresponding to a weight hourly spacevelocity (WHSV) of 7.00 h⁻¹), constituting a mass flow rate of 296 g/h,together with hydrogen at a flow rate of 40.0 Nl/h. The hydrogen/dienemolar ratio was 4.10. The total pressure was 30 bar and the inlettemperature was varied from about 45° C. to about 120° C.

The composition of the effluent of the reactor was analysed over thevarying inlet temperatures and the results are summarised in FIG. 1.

From FIG. 1 it will be seen that for the aromatics content of theeffluent, this was substantially unchanged as compared to the aromaticscontent of the feedstock. The olefins content was increased in theeffluent as compared to that in the feedstock. However, the olefinscontent tended to decrease with increasing inlet temperature up to 120°C. For inlet temperatures of from about 45 to 80° C., the olefinscontent was about 17 wt %, decreasing gradually to about 14 wt % at aninlet temperature of 120° C. For the paraffins content, this wasincreased in the effluent as compared to the paraffins content of thefeedstock. The paraffins content gradually increased with increasinginlet temperature. Thus at an inlet temperature of about 45° C. theparaffins content was about 27 wt %, increasing to a paraffins contentof about 33 wt % at an inlet temperature of about 120° C. Mostsignificantly, the dienes content of the effluent was significantlyreduced as compared to that of the feedstock, and the dienes content ofthe effluent tended to decrease yet further with increasing inlettemperature. Thus at inlet temperatures of about 45° C., the dienescontent was about 2 wt %, significantly less than the original dienescontent of about 12 wt % and the dienes content of the effluentdecreased to about 0.25 wt % at an inlet temperature of about 120° C.

The significant decrease in the dienes content of the effluent ascompared to that of the feedstock, with a corresponding smaller increasein the paraffins and olefins content, and with the aromatics contentbeing substantially unchanged, indicates the effectiveness of theselected hydrogenation catalyst of the present invention. Thus thecatalyst is very active for the hydrogenation of dienes, and a goodselectivity for olefins is maintained.

TABLE 1 Composition SiO₂ 49.0 wt % CaO 42.9 wt % Al₂O₃ 0.2 wt % MgO 0.3wt % Fe₂O₃ 1.1 wt % Na₂O 0.2 wt % K₂O 0.2 wt % Loss on Ignition 6.1 wt %Specific area (BET) 50 m²/g Bulk Density 90 g/l Average particle size 45μm

TABLE 2 Feedstock Composition Paraffins 24.46 wt % Olefins 10.91 wt %Dienes 12.20 wt % Aromatics 52.43 wt % Diene Value [gram I₂/100 gram]18.21 Sulphur 94 wppm Density 0.802 g/ml

1. A process for hydrogenating unsaturations in a petrochemicalfeedstock, the process comprising contacting the petrochemicalfeedstock, including at least one component having unsaturations, andhydrogen with a catalyst comprising at least one Group Ia, Ib, IIb, VIb,VIb or VIII metal impregnated on a support of a crystalline calciumsilicate having a surface area of at least 30 m²/g, the support being inthe form of substantially spherical particles having a mean diameter offrom 10 to 200 microns and comprising pores in the particles having adiameter of from 100 to 2000 Angstroms, at a temperature of from 0 to550° C. and a pressure of from 3 to 150 barg.
 2. A process according toclaim 1 wherein the calcium silicate has the chemical compositionCa₆Si₆O₁₇(OH)₂.
 3. A process according to claim 1 wherein the supporthas a basicity corresponding to a pH of greater than 7.5.
 4. A processaccording to claim 1 wherein said catalyst comprises palladiumimpregnated onto the support in an amount of from 0.01 to 10 wt. % basedon a weight of the supported catalyst.
 5. A process according to claim 1wherein the petrochemical feedstock is passed over the catalyst at aliquid hourly space velocity of from 0.1 to 100 h⁻¹.
 6. A processaccording to claim 1 wherein the molar ratio of hydrogen to the at leastone component having unsaturations which is hydrogenated is from 0.7 to200.
 7. A process according to claim 1 wherein the at least onecomponent having unsaturations which is hydrogenated comprises butadienein a C4 stream.
 8. A process according to claim 1 wherein the at leastone component having unsaturations which is hydrogenated comprises atleast one of vinyl acetylene and ethyl acetylene in a C4 stream. 9.(canceled)
 10. A process according to claim 1 wherein the at least onecomponent having unsaturations which is hydrogenated comprises at leastone of methyl acetylene and propadiene in a C3 stream.
 11. (canceled)12. A process according to claim 1 wherein the at least one componenthaving unsaturations which is hydrogenated comprises at least one of adiolefin and an unsaturated aromatic in pyrolysis gasoline. 13.(canceled)
 14. (canceled)
 15. A process according to claim 1 wherein theat least one component having unsaturations which is hydrogenatedcomprises phenyl acetylene in a styrene stream.
 16. A process accordingto claim 1 wherein the at least one component having unsaturations whichis hydrogenated comprises olefins in an aromatic rich fraction. 17.(canceled)
 18. (canceled)
 19. A process according to claim 1 wherein thesupport is in the form of substantially spherical particles having amean diameter of from 10 to 200 microns and comprising pores in theparticles having a diameter of from 100 to 2000 Angstroms. 20.(canceled)
 21. A process according to claim 1 wherein the support has asurface area of from 30 to 200 m²/g.
 22. (canceled)
 23. A processaccording to claim 1 wherein the petrochemical feedstock is passed overthe catalyst at a liquid hourly space velocity of from 1 to 100 h⁻¹, aninlet temperature of from 0 to 250° C. and a pressure of from 3 to 50barg.
 24. (canceled)
 25. A process according to claim 1 wherein the atleast one component having unsaturations which is hydrogenated isselected fi-n the group consisting of butadiene in a C4 stream; at leastone of methyl acetylene and propadiene in a C3 stream; at least one of adiolefin and an unsaturated aromatic in pyrolysis gasoline; at least oneof a diene and an alkyne in pyrolysis gasoline; phenyl acetylene in astyrene stream; and an alpha-olefin in an aromatic rich fraction. 26-29.(canceled)
 30. (canceled)
 31. A process for dehydrogenating or reforminga hydrocarbon feedstock comprising contacting the catalyst of claim 1with said feedstock under conditions effective to dehydrogenate orreform said feedstock. 32-33. (canceled)
 34. A catalyst supportcomprising a crystalline calcium silicate of molecular formula6CaO.6SiO₂.H₂O having a surface area of at least 30 m²/g, the Supportbeing in the form of substantially spherical particles having a meandiameter of from 10 to 200 microns and pores in the particles having adiameter of from 100 to 2000 Angstroms, wherein a catalyst is producedby impregnating at least one Group Ia, Ib, IIb, VIb, VIIb or VIII metalon the calcium silicate support.
 35. (canceled)
 36. (canceled) 37.(canceled)
 38. (canceled)
 39. (canceled)
 40. The process according toclaim 1 wherein the at least one component comprises at least one vinylacetylene and an ethyl acetylene in a C₄ stream.
 41. The processaccording to claim 1 wherein the C₄ stream is from a FCC unit, avisbreaker, a coker, or a steam cracker.
 42. (canceled)
 43. The processaccording to claim 25 wherein the C₃ stream is from a steam crackingunit.
 44. (canceled)
 45. (canceled)
 46. The process according to claim12 wherein the pyrolysis gasoline is from a steam cracking unit, a cokerunit, a visbreaker, or comprises a light cracked naphtha from a FCCunit.
 47. (canceled)
 48. The process according to claim 36 wherein theat least one component comprises an olefin in an aromatic rich fraction.49. (canceled)
 50. (canceled)
 51. (canceled)
 52. (canceled) 53.(canceled)
 54. (canceled)
 55. A process for dehydrogenating or reforminga hydrocarbon feedstock using the catalyst of claim 1 comprisingcontacting said catalyst with said feedstock under conditions effectiveto dehydrogenate or reform said feedstock.